Process for treating loaded extractant from purification of phosphoric acid by extraction and recovering nutrients

ABSTRACT

In a process for purifying an aqueous phosphoric acid containing up to about 55% P 2  O 5  by weight and dissolved magnesium, iron and aluminum ions, wherein (a) said phosphoric acid is contacted with an organic extractant containing a water immiscible organic sulfonic acid to form: a loaded organic extractant phase containing the organic sulfonic acid, extracted ionic metallic impurities and extracted P 2  O 5  values, and a purified aqueous phosphoric acid; (b) the purified aqueous phosphoric acid phase is separated from the loaded organic extractant phase; (c) the separated loaded organic extractant phase is contacted with a solution phase containing calcium ions (e.g., calcium nitrate), under conditions whereby at least some of said extracted ionic metallic impurities (e.g., Mg +2 , Al +3 , Fe +3 ) are replaced by calcium and are transferred to said solution phase; and (d) the solution phase is separated from the contacted extractant phase; the improvement wherein the separated solution phase from step (d) is made basic, preferably with lime, to cause a solid precipitate to form, said solid precipitate comprising at least some of the metallic impurities and/or P 2  O 5  values from phosphoric acid which was entrained in the organic extractant. The separated solid precipitate is useful as a nutrient (e.g., as a fertilizer or as an animal feed supplement). Where sulfuric acid is used for the acidification, solid calcium sulfate forms, which can be separated from the &#34;used&#34; acid. The sulfuric acid can be strengthened with fresh, more concentrated, acid and recycled.

CROSS REFERENCE TO RELATED APPLICATIONS AND PATENTS

This application is related to the processes for purification ofphosphoric acid by extraction of U.S. Pat. Nos. 3,694,153 to Williamsand Stern, 4,053,564 to Bradford and Ore, 4,082,836 to Ore and pendingU.S. applications Ser. No. 840,791 filed Oct. 11, 1977 to Bradford andOre titled "Uranium Recovery from Pre-Treated Phosphoric Acid" and Ser.No. 676,559 filed Apr. 13, 1976 of Ore titled "Process for theProduction of High Purity Phosphoric Acid from Phosphate Rock andPebble". The entire disclosure of all of these patents and applicationsis hereby incorporated herein by reference.

BACKGROUND OF THE INVENTION

Superphosphoric acid (SPA) is a condensation product of orthophosphoricacid. The minimum P₂ O₅ content of SPA is about 65% by weight, themaximum is greater than 100%. The P₂ O₅ content of most commercialgrades of SPA is from 72% to about 85%.

Wet process phosphoric acid can be converted to SPA by means ofconventional evaporation and dehydration techniques. SPA formed fromunpurified wet process phosphoric acid is usually unstable due to ionicmetallic impurities present therein. The impurities cause an increase inthe viscosity of the SPA to such an extent that a semi-solid orglass-like material results. Further processing of such materials isextremely difficult.

On the other hand, purified wet process phosphoric acid yields a lowviscosity SPA which can be readily handled (e.g. see U.S. Pat. Nos.3,044,851 and 3,192,013) Williams et al, 3,694,153 and 4,053,564 ofBradford et al, removed ionic metallic impurities from phosphoric acidemploying water immiscible organic sulfonic acids. Using solventextraction techniques, wet process phosphoric acid was purified byremoving the ionic metallic impurities therefrom.

While removing the impurities, P₂ O₅ values are entrained or coextractedwith the metallic impurities into the organic phase. These P₂ O₅ valuesare usually lost when the water immiscible organic sulfonic acid isregenerated with a mineral acid because the mineral acid regenerationsolution containing the ionic metallic impurities and the P₂ O₅ valuesis sent to waste disposal.

U.S. Pat. No. 4,082,836 of Ore recovers P₂ O₅ values in the organicextract by contacting the organic extractant with a wash phasecomprising water or dilute phosphoric acid to form a loaded wash phaseenriched with P₂ O₅ values extracted from the organic extractant. Theloaded wash phase is then separated from the organic extractant phase.The washed organic extractant can then be regenerated on treatment witha mineral acid and be recycled into the process. In the Ore' process(and in those of U.S. Pat. Nos. 3,694,153 and 4,053,564) the sulfuricacid consumption to regenerate the organic extractant is about 10 timesstoichiometric for removal of metal ions. Also, the waste acid must beneutralized (as with lime) before disposal. In the process of thepresent invention, sulfuric acid and lime concentration is much lessthan that half that in these prior art processes.

Aqueous phosphoric acid phase denotes a solution of phosphoric acidcontaining from about 1% to about 55% P₂ O₅ by weigh which will be, isbeing, or has been subjected to solvent extraction. Terms such asphosphoric acid phase, crude phosphoric acid feed, unpurified phosphoricacid solution and purified phosphoric acid can be used to indicate anaqueous phosphoric acid phase.

Organic extractant phase denotes a solution in which at least oneorganic sulfonic acid is dissolved in an organic solvent and preferablycontaining a water immiscible organic phosphate or phosphonate and canalso contain a water immiscible aliphatic alcohol. Extractant phase,organic phase, barren organic phase, loaded organic extractant phase,and regenerated organic extractant phase have been used to denote theorganic extractant phase at some step in the solvent extraction process.

Wash phase denotes the aqueous phase (water or a dilute solution ofphosphoric acid containing from about 10 grams/liter to about 300grams/liter of P₂ O₅) which is used to remove P₂ O₅ values from theorganic extractant phase; said P₂ O₅ values being coextracted with metalimpurities into the organic extractant phase during the extracting stepof solvent extraction process. Terms used in denote the wash phase aredilute aqueous phosphoric acid, dilute phosphoric acid, dilutephosphoric acid phase, dilute phosphoric acid stripping phase, waterwash phase, loaded wash phase, water wash liquor enriched phosphoricacid solution and enriched phosphoric acid phase.

The mineral acid stripping phase is the solution of an inorganic acidsuch as sulfuric acid, nitric acid or hydrochloric acid which is used toregenerate the organic sulfonic acid present in the organic extractantphase. Sulfuric acid stripping solution, sulfuric acid stripping phase,fresh sulfuric acid stripping solution, stripping agent, sulfuric acidphase, waste sulfuric acid stripping, waste sulfuric acid strippingsolution, waste solution and waste stripping agent are used to denotethe mineral acid stripping phase.

A mineral acid is an inorganic acid having an ionization constant equalto 10⁻³ or greater.

SUMMARY OF THE INVENTION

The present invention is directed to a process for purifying phosphoricacid using solvent extraction techniques. An unpurified aqueousphosphoric acid phase typically containing from 1% to 55% P₂ O₅ byweight and dissolved ionic metallic impurities such as calcium,magnesium, iron (II & III) and aluminum is throughly mixed with anorganic extractant phase containing at least one organic sulfonic acidin the H⁺ form. As a result of the thorough mixing, an organicextractant phase containing at least one organic sulfonic acid,extracted ionic metallic impurities and extracted P₂ O₅ values is formedand separated from the purified aqueous phosphoric acid phase.

If phosphoric acid product is desired, the organic extractant phase iscontacted with a wash phase (water or dilute phosphoric acid, preferablycontaining from about 10 grams/liter to about 300 grams/liter of P₂ O₅)to form a loaded wash phase enriched with P₂ O₅ values extracted fromthe organic extractant phase. The loaded wash phase is separated fromthe organic extractant phase. The organic extractant phase containingextracted ionic metallic impurities, either washed or unwashed, iscontacted with a solution of a calcium salt (e.g. calcium chloride,calcium nitrate), whereby ionic metallic impurities, comprising ions ofmagnesium and of aluminum, are removed from the extractant phase andreplaced therein by calcium ions, followed by contact with a base (e.g.lime) to cause precipitation of a solid phase containing compounds ofmagnesium and aluminum and P₂ O₅ values.

If desired, the compounds with P₂ O₅ values can be separatelyprecipitated below about pH10 from the compounds of magnesium (at aboveabout pH10). These precipitates have a low fluoride content and areuseful as an animal feed supplement or as a fertilizer.

The organic extractant phase can be regenerated on treatment with amineral acid and recylced into the process.

This invention also can provide a solution to an existing effluentproblem (e.g., waste acid from regeneration) in the solvent extractionof wet process phosphoric acid and also can provide a means of producingan animal feed supplement.

In the process of Ore', U.S. Pat. No. 4,082,836, where P₂ O₅ isrecovered in a wash step, as much as 5% of the P₂ O₅ in the impure acidfeed can be lost in the waste sulfuric acid. In the present process,this P₂ O₅ can be recovered as phosphates (of calcium and perhapsaluminum).

In a preferred embodiment, the magnesium in loaded extractant (which canbe washed) is removed from the organic phase by a contact with aqueouscalcium nitrate solution. Magnesium attached to the active sulfonic acid(DNSA) in the extractant is replaced by calcium ion. Magnesium ion istransferred to the aqueous phase. Thus a magnesium-loaded extractantafter substitution becomes predominantly loaded with calcium ion. Thisis represented by the following reaction: ##EQU1## The followingadditional reaction also takes place simultaneously: ##EQU2## Theorganic phase (not containing Ca²) is then stripped with a solution ofsulfuric acid, whereby the extractant is regenerated and calcium sulfate(gypsum) precipitates in the aqueous phase. Gypsum is filtered and thesulfuric acid recycled.

The aqueous stream now containing nitric acid, magnesium nitrate, andexcess calcium nitrate is treated with lime (e.g. PH of 10.5) wherebymagnesium is precipitated out. The clear liquor from this operation isrecycled to the substitution stage. Aluminum removed from the extractantalso precipitates at this step.

In the embodyment described above, there is no liquid waste stream.

THE DRAWING

The accompanying drawing,

FIG. 1, is a schematic of the present solvent extraction process.

FIG. 2 is a schematic description of the process of Example 1.

FIG. 3 is a schematic description of the process of Example 2.

DETAILED DESCRIPTION

The present process is directed to the purification of phosphoric acidby means of solvent extraction. The process is comprised of anextraction step, a calcium solution (e.g., calcium nitrate contact)step, a neutralization (e.g., lime contact) step and a mineral acidstripping step.

When the solution contains calcium chloride or calcium nitrate, thepreferred concentration is in a range of about 1% to saturated at thecontact temperature, typically 5 to 15% by weight.

Where more than one Calcium solution contact stage is used in acocurrent operation, the later stages are preferably a lower calciumconcentration than the earlier stages. This is because the loadedorganic is usually far from fully exchanged with metals and high calciumconcentration in multiple stages can cause excess calcium substitutionin the organic phase (for hydrogen rather than metals).

Where it is desired to produce additional, highly purified phosphoricacid, a wash phase stripping step can be used prior to the contact ofthe extractant with the calcium solution.

The extraction step comprises mixing an aqueous phosphoric acid phasewith an organic extractant phase. The aqueous phosphoric acid phase,typically containing from 1% to 55% P₂ O₅ by weight and ionic metallicimpurities such as Ca⁺², Mg⁺², Al⁺³, and Fe⁺²,+3, is thoroughly mixedwith an organic extractant phase (also referred to as the barren organicphase) which contains at least one water immiscible organic sulfonicacid in its H⁺ form. Ionic metallic impurities are extracted from theaqueous phosphoric acid phase into the organic extractant phase. Aftermixing, the purified aqueous phosphoric acid phase separates from theorganic extractant phase, which now contains ionic metallic impuritiesand some P₂ O₅ values extracted from the aqueous phosphoric acid phase.Separation, preferably occurs in a mixer-settler cell, due to densitydifferences and the immiscibility of the two phases. After separation,removal is effected by pumping, decanting, gravity flow or the like.

The separated organic extractant phase containing the ionic metallicimpurities and P₂ O₅ values is mixed with aqueous calcium nitrate,whereby ionic metallic impurities comprising ions of magnesium andaluminum in the extractant phase are replaced by calcium. The aqueouscalcium nitrate effluent extractant phase is then treated with a base,preferably lime, to cause precipitation of solids (containing compoundsof magnesium and aluminum and P₂ O₅ values). These solids can be usedper se as an animal feed supplement, but more preferably are blendedwith additional nutrients, such as compounds containing calcium and/orphosphorous (e.g., low fluorine content dicalcium phosphate or"defluorinated" phosphate rock.

Where a lower P₂ O₅ content solid is desired, or where it is desired toproduce additional phosphoric acid prior to contact with the calciumnitrate, the extractant containing metallic impurities is contacted witha wash phase, which can be water or, more preferably, a dilutephosphoric acid phase containing from about 10 grams/liter to about 300grams/liter of P₂ O₅. The wash phase extracts the P₂ O₅ values over theionic metallic impurities from the organic extractant phase. The washphase enriched with P₂ O₅ values extracted from the organic extractantphase is separated from the organic extractant phase. Separation occursdue to density differences and the mutual immiscibility of the phases ina mixer-settler cell.

The separated wash phase can be used as a source of purified phosphoricacid, or it can be mixed with the aqueous phosphoric acid phase which isthe feed for the extraction step or it can be diluted with water to formthe wash phase, that is the dilute phosphoric acid phase, used torecover P₂ O₅ values from the organic extractant phase.

The organic extractant phase after separation from the soluble calciumcontact stage can be treated with a mineral acid stripping phasepreferably containing sulfuric, nitric or hydrochloric acid, toregenerate the sulfonic acid extractant to its H⁺ form.

The most preferred mineral acid is sulfuric acid, because the calcium inthe extractant, which is replaced by hydrogen from the acid, combineswith the sulfate ions to form solid calcium sulfate (gypsum) which iseasily separated by filtration. This permits the acid phase to bestrengthened by adding concentrated acid and be recycled in a closedloop system.

When nitric acid is used for the stripping, the separated "used" acidcontaining calcium nitrate can be recycled or can be contacted with limeor limestone to produce calcium nitrate for use in removing magnesium,etc. from loaded extractant. If the nitric acid is recycled, asidestream can be taken off and converted to calcium nitrate and fresh,concentrated acid added to maintain calcium exchange capability.

Similarly, if hydrochloric acid is used and recycled, a sidestream canbe contacted with lime or limestone and the resulting calcium chloridesolution can be used in removing magnesium, etc. from loaded extractant.Separation of the organic extractant phase from the mineral acidstripping phase in a mixer-settler cell results from differences indensity and mutual immiscibility of the two phases. Separation iseffected by pumping, decanting, gravity flow or the like. Thereafter theregenerated organic extractant phase can be recycled into the extractionstep described above.

The process is preferably run at a temperature of from about 50° F. toabout 65° C. to reduce P₂ O₅ loss to the organic extractant phase and toincrease the speed of separation of the organic phases from the aqueousphases.

The aqueous phosphoric acid phase can be wet process phosphoric acidproduced by the dihydrate, the hemihydrate or the anhydrite process; itmay be a dilute (1.0% to 5% P₂ O₅ by weight) solution produced by theacidulation of P₂ O₅ values in slimes produced in the benefication ofphosphate rock, or it may be a waste stream or phosphoric acid which cancontain, in addition to the impurities mentioned, also nickel, copper,zinc, chromium and the like. The impure phosphoric acid which is treatedwill usually contain from about 26% to about 58% P₂ O₅ by weight and canbe produced by the wet process or by a reduction process, such as thoseof U.S. Pat. Nos. 1,422,699, 3,052,523; 3,056,659; 3,235,330; 3,241,917;and 3,479,138.

The temperature of the aqueous phosphoric acid phase entering thesolvent extraction system is from about 20° C. to about 77° C.,preferably from 55° C. to about 70° C.

The organic extractant phase consists of one or more water immiscibleorganic sulfonic acids. Preferably the water immiscible organic sulfonicacids are dissolved in a water immiscible organic solvent.

Water immiscible organic sulfonic acids employed in accordance with thepractice of this invention includes aryl-, alkylaryl-, polyalkylaryl-,alkanoylaryl-, and polyalkanoylaryl-sulfonic acids.

The water immiscible organic sulfonic acids contain at least 12 carbonatoms and preferably about 14 to about 30. An especially preferred groupof water immiscible organic sulfonic acids for the present processinclude dinonylnaphthalene sulfonic acid,5-dodecanoyl-2-chlorobenzenesulfonic acid,5-nonyl-2-ethoxybenzenesulfonic acid,3,5-di-ti-octyl-2(alphacarboxymethoxy) -benzenesulfonic acid and3,5-di-t-octyl-2-hydroxybenezenesulfonic acid.

The water immiscible organic sulfonic acid can be dissolved in the waterimmiscible organic solvent in concentrations of about 0.05 molar toabout 3.0 molar of the water immiscible organic sulfonic acid. Thepreferred concentration range of the sulfonic acid in the organicsolvent is from about 0.2 molar to about 0.5 molar.

The preferred solvents are saturated hydrocarbons having a boiling pointbetween 120° C. and 230° C., flash points between 15° C. and 70° C. anddensities lower than the density of the phosphoric acid.

Illustrative, but nowise limiting, of the water immiscible organicsolvents for the organic sulfonic acids employed in the present processare kerosene, mineral spirits, benzene, alkyl-substituted aromaticcompounds and alkyl-substituted haloaromatic compounds, such as xylene,toluene, ethylbenzene, chlorotoluene and the like, nitrobenzene, carbontetrachloride, chloroform, methylene chloride, trichloroethylene,isooctane, heptane and the like. As is well known, materials such askerosene and mineral spirits are mixtures of compounds.

When using solvent extraction techniques to effect transfer of solutesfrom one phase to another phase, there can be a problem with emulsionformation. This problem can be minimized by employing water immisciblealiphatic alcohols, water immiscible acid phosphates or phosphonates andmixtures thereof.

Water immiscible aliphatic alcohols can be added to the organic phaseand in addition to aiding in the separation of the organic phase fromthe aqueous phase, they tend to increase the solubility of the waterimmiscible organic sulfonic acid in the organic phase. Water immisciblealiphatic alcohols are preferred, straight chain or branched chainaliphatic alcohols being especially preferred. Water immisciblealiphatic alcohols which can be employed in the present process areoctanol, decanol and isodecanol. Isodecanol is especially preferred.

For the present process, it has been found that the water immisciblealiphatic alcohol can be used in concentrations from 5% to about 20% byweight, based on the weight of the organic extractant phase. Aconcentration of 7% to about 12% by weight, based on the organicextractant phase, is preferred.

More preferred than a water immiscible alcohol is at least one waterimmiscible organic acid phosphate, phosphonate or half ester or mixturesthereof with a water insoluble alcohol. These organic acid phosphatesare described, for example, in U.S. Pat. No. 4,053,564 and includemixtures of mono-and di-organo substituted phosphoric acid compounds.Typically such a mixture can contain in the range of about 0.1 to about60.0% mono-substituted and about 97% to about 35% by weightdisubstituted organic phosphates. The organic phosphates include thefollowing (mono or di substituted) phosphates: amyl acid phosphate,isooctyl acid phosphate, 2-ethylhexyl acid phosphate, decyl acidphosphate, lauryl acid phosphate, stearyl acid phosphate, phenyl acidphosphate and octylphenyl acid phosphate, cyclohexyl acid phosphate,cyclohexyl phenyl acid phosphate, 1-(5-hexynyl) acid phosphate,1-(5-hexenyl)-phenyl acid phosphate, 1(6-chlorohexyl) acid phosphate,1-(6-hydroxyhexyl) acid phosphate and 1-(6-methoxy-hexyl) acidphosphate. When aryl-, alkylaryl or alicyclicaryl-acid phosphates areemployed, the aromatic ring can be substituted with a hydroxy, an alkoxyor a halo group.

Water immiscible organic phosphonates can also be used in thisinvention. The half-ester can be used. The formula for the phosphonatesis R-P(O) (OH) (OR°) in which R and/or R° can be alkyl, alkenyl,alicyclic, aryl, alkenylaryl, alkylaryl, alicyclicaryl or heterocyclic.The organic substituent groups, R and R°, may be substituted with analkoxy, a hydroxy or a halo group; said R groups may be the same ordifferent. R and/or R° should contain from 4 to about 18 carbon atoms.

The water immiscible organic sulfonic acid and the water immiscibleorganic acid phosphate can be dissolved in water immiscible organicsolvents. The preferred solvents are saturated hydrocarbons having aboiling point between 120° and 230° C., flash points between 15° and 80°C. and densities lower than the density of the phosphoric acid. Eitherpure hydrocarbons or mixtures of hydrocarbons may be used.

Illustrative, but nowise limiting, of the water immiscible organicsolvent for the water immiscible organic sulfonic acids and the waterimmiscible organic acid phosphates employed in the present process arekerosene, mineral spirits, benzene, naphtha, xylene, toluene,nitrobenzene, carbon tetrachloride, chloroform, methylene chloride,trichloroethylene, isooctane, heptane and the like.

The water immiscible organic sulfonic acid can be dissolved in the waterimmiscible organic solvent in concentrations of about 0.05 to about 3.0molar or more of the water immiscible organic sulfonic acid. Thepreferred concentration range of the water immiscible organic sulfonicacid in the water immiscible organic solvent is from about 0.2 to about0.5 molar.

The water immiscible organic acid phosphate can be dissolved in thewater immiscible organic solvent in concentrations of 0.05 to about 1.0molar or more of the water immiscible organic acid phosphate. Thepreferred concentration is about 0.15 to about 0.65 molar of the waterimmiscible organic acid phosphate.

The organic extractant phase enters the solvent extraction system at atemperature of about 20° C. to about 60° C., preferably from about 30°C. to about 58° C.

The phosphoric acid phase at a temperature of from 20° to 77° C. ismixed with the organic extractant phase at a temperature of about 20° toabout 60° C; the temperature of the resulting mixture is from about 20°to about 65° C.

The volume ratio of the organic extractant phase to the aqueousphosphoric acid phase in the present process is from about 1 to 1 toabout 15 to 1, preferably from about 2 to 1 to about 12 to 1. Anespecially preferred volume ratio is from about 2 to 1 to about 8 to 1of the organic extractant phase to the aqueous phosphoric acid phase.

The organic extractant phase is virogously mixed with the aqueousphosphoric acid phase at least once during the purification process.However, the organic extractant phase can be contacted with the aqeuousphase a number of times during the process. The number of separatecontacts can be from 1 to about to 20 or more; preferably at least 3contacts are utilized in the present invention.

The organic extractant phase contacts the aqueous phospheric acid phaseand is separated therefrom. Separation in a mixer-settler cell resultsfrom difference in densities and the mutual immiscibility of the phases.

The organic extractant phase extracts P₂ O₅ values as well as ionicmetallic impurities from the aqueous phosphoric acid phase. The P₂ O₅values can range from 3% to about 20% of the P₂ O₅ initially present inthe aqueous phosphoric acid phase. Recovery of these P₂ O₅ values fromthe organic extractant phase is desirable. The present invention permitsrecovery of these P₂ O₅ values as a solid e.g., comprising, calciumphosphates or, with a wash step, as a solid and a phosphoric acidsolution.

Where a wash phase is used, the organic extractant phase, containing theionic metallic impurities and P₂ O₅ values extracted from the aqueousphosphoric acid phase, is virgorously mixed with the wash phase. Jetmixers, line mixers, centrifugal pumps, mechanical agitators and thelike can be used to effect the mixing. Separation in a mixer-settlercell results from differences in densities and the mutual immiscibilityof the two phases. Separation is effected by pumping, decanting, gravityflow or the like. Separation of the two phases gives an organicextractant phase and a loaded wash phase. The loaded wash phase hasremoved P₂ O₅ values from the organic phase. The temperature maintainedin the wash phase is from about 20° C. to about 70° C., preferably fromabout 55° C. to about 65° C.

The volume ratios of the organic extractant phase to the calciumsolution (e.g. calcium nitrate) phases can be from 1 to 1 to about 10 to1, preferably from about 2 to 1 to about 4 to 1.

After thorough mixing of the calcium nitrate phase with the organicextractant phase at a temperature of from about 20° to about 65° C., thewash phase and the organic extractant phase are separated. Separationpreferably occurs in a mixer-settler cell due to the differences indensity and the immiscibility of the phases. The phases are separated bypumping, decanting gravity flow or the like.

The organic extractant phase separated from the loaded calcium nitratephase can be regenerated by treatment with a mineral acid phase. Ionicmetallic impurities and residual P₂ O₅ values are removed from theorganic extractant phase by means of a mineral acid stripping treatmentand the water immiscible organic sulfonic acid is regenerated to its H⁺form.

The mixing of the mineral acid stripping phase and the organicextractant phase is achieved with a jet mixer, line mixer, centrifugalmixer, a mechanical agitator or the like. The temperature of the mixtureis from about 20° to about 65° C. Separation of the two phases occurs ina mixer-settler cell due to immiscibility of the two phases and adifference in specific gravities. The phases are further separated bypumping, decanting, gravity flow or the like.

The volume ratio of the organic extractant phase to the mineral acidstripping phase can be in the range from about 0.5 to 1 to about 15 ormore to 1, preferably 1 to 1 to about 5 to 1, at a ratio less than 1:1,an aqueous continuous phase can form at low acid strengths. The organicextractant phase is contacted with the mineral acid stripping phase atleast once, preferably from 1 to about 10 times and most preferably onlyonce.

The mineral acid used in stripping the organic extractant phase can besulfuric, nitric, or hydrochloric acid; sulfuric acid being greatlypreferred because insoluble calcium sulphate can be readily separatedfrom the liquid (acid and organic extractant). The concentration of thehydrochloric acid and the nitric acid can be from about 2% acid to about60% (or more) acid by weight. The concentration of the sulfuric acid inthe mineral acid stripping phase can be from about 2% to about 50.0% H₂SO₄ by weight (the upper limit being determined by excess viscosity),preferably from 15% to 30% H₂ SO₄ by weight. The mineral acid strippingphase is maintained at about 20° C. to about 77° C., preferably from 60°C. to 71° C. The temperature of the mineral acid stripping phase isachieved by utilizing the heat of dilution of the mineral acid withwater and by using hot steam condensate as the diluent for the mineralacid.

The hot mineral acid stripping phase (illustrated hereinafter bysulfuric acid) removes the calcium ions and any other residual extractedionic metallic impurities and residual P₂ O₅ values from the organicextractant phase. The mineral acid stripping phase after separation (asdescribed) from the regenerated organic extractant phase containsprincipally calcium sulfate solids and unused sulfuric acid. It can alsocontain small amounts of phosphoric acid and the sulfate salts of otherionic metallic impurities. When the initial concentration of thesulfuric acid in the mineral acid stripping phase is from 10% to 30% H₂SO₄ by weight, considerable amounts of sulfuric acid are present in thewaste mineral acid stripping phase which can be utilized by mixing itwith fresh mineral acid stripping phase. Alternately, all or a portionof waste mineral acid stripping phase can be sent to waste disposalwhere the sulfuric acid in the mineral acid stripping phase is reactedwith lime or the like to form gypsum. Preferably at least astoichiometric amount of sulfuric acid is used to regenerate the organicsulfonic acid in the organic extractant phase. The regenerated organicextractant phase, having the water immiscible organic sulfonic acidcontained therein converted back to its H⁺ form, is transported to theorganic extractant phase storage area (barren organic phase) for use inthe above extraction process.

U.S. Pat. No. 4,082,836 illustrates a mixer-settler typical of thoseused in the present process. The mixer-settler is a rectangularcontainer, the length being greater than the width or depth. The topportion is open, although during operation it can be covered to reducesolvent loss, decrease fire hazard and reduce the introduction ofcontamination therein. The mixer-settler is divided into severalsections. A mixing chamber containing an agitator and an organicextractant phase inlet and an aqueous phosphoric acid phase inlet arelocated at one end of the mixer-settler. The inlets are located in thebottom portion of the mixing chamber in order to achieve good contactbetween the organic and aqueous phases. The organic extractantphase-aqueous phase mixture passes over a weir when mixing chamber isfilled into a settling chamber. Here the phases separate into an organicextractant phase and an aqueous phosphoric acid phase. At times, a smallemulsion band, several inches thick can occur at the interface. This canbe reduced by prior contact of the phosphoric acid with activatedcarbon, or the other means suggested in Ser. No. 840,791. Recoverychambers are separated by a wall. When the settling chamber becomesfilled, the organic extractant phase passes over a weir into an organicextractant phase recovery chamber. At the same time, the aqueous phasepasses through an adjustable standpipe which controls the position ofthe interface into the aqueous phase recovery chamber. Once in therecovery chamber, the organic extractant phase pases through an outletand is sent either storage or to another mixer-settler cell for furthertreatment. Aqueous phophoric acid phase passes through outlet and issent to storage as purified phosphoric acid or to another mixer-settlercell.

In a continuous process, several mixer-settler cells can be arranged inseries with the organic extractant phase moving through the cells inseries counter-currently to the movement of the aqueous phosphoric acidphase through the cells.

The FIGURE is schematic of one preferred embodyment of the process forthe purification of phosphoric acid.

Aqueous phosphoric acid phase 38 prepared by the dihydrate, thehemihydrate or the anhydrite process enters mixer-settler cell 35. Theincoming aqueous phosphoric acid phase 38, at a temperature of about 20°C. to about 77° C., is mixed with an organic extractant phase 36, at atemperature of about 20° to about 60° C., which contains considerablequantities of ionic metallic impurities and P₂ O₅ values as the resultof having passed through mixer-settler cells 31 and 33, which aresimilar to the cell 2 illustrated in FIG. 1. The temperature of themixture is about 50° to about 65° C. The organic extractant phase 36 andthe aqueous phosphoric acid phase 38 so mixed passes into the settlingchamber (not shown) of mixer-settler cell 35 in which the two phasesseparate. The aqueous phosphoric acid phase 40 passes into mixer-settlercell 33, where the mixing, settling and separation steps described aboveare repeated. The organic extractant phase 36 contacting the aqueousphosphoric acid phase 40 in mixer-settler cell 30 contains lessextracted ionic metallic impurities than the organic extractant phaseentering cell 35 because it has contacted the aqueous phosphoric acidphase in only one cell.

On exiting cell 33, the aqueous phosphoric acid phase 40 is sent tomixer-settler cell 31, and is mixed with organic extractant phase 34entering cell 31. The organic extractant phase 34 enters the extractionsection at a temperature of about 20° C. to about 60° C., preferablyfrom about 50° C. to about 58° C. The temperature of the resultingmixture is about 50° C. to about 65° C.

The mixing, settling and separation process is repeated in cell 31. Theaqueous phosphoric acid phase 40 exiting cell 31 is purified to such anextent that it is suitable for SPA production and passes to purifiedphosphoric acid storage (not shown).

The organic extractant phase 36 extracts ionic metallic impurities fromaqueous phosphoric acid phase 40 in cells 31, 33, and 35; in addition,P₂ O₅ values are coextracted into the organic extractant phae 36. Theorganic extractant phase 36 containing ionic metallic impurities and P₂O₅ values, after leaving the extraction section, enters mixer-settlercell 37 of the wash section.

If it is desired to maximize the yield of phosphoric acid rather thanproducing low fluoride solids (suitable as an animal nutrient), washphase 42 can be mixed with phosphoric acid at T-junction 44. Theresulting dilute phosphoric wash phase 43a passes into cell 37. Theorganic extractant phase 36 is vigorously mixed with the wash phase 43ain cell 37. P₂ O₅ values are extracted from the organic extractant phase36 into the wash phase 43a. After settling, the organic and aqueousphases are separated. The loaded wash phase is divided into streams 43and 47. The loaded wash phase is diluted with incoming water 42, atT-junction 44 and recycled back into cell 37. The stream 47 is mixedwith aqueous phosphoric acid phase 38 (not shown) and re-entersmixer-settler cell 35.

The organic extractant phase 36, exiting cell 37, contains ionicmetallic impurities and residual P₂ O₅ values and is treated inmixer-settler cells 51 and 54 with a calcium solution phase (e.g.,aqueous calcium nitrate) 52. Cells 51 and 54 can be similar in structureto the cell illustrated in U.S. Pat. No. 4,082,836.

The metal stripping section preferably comprises a treatment withcalcium solution, a neutralization step and a mineral acid strippingphase. In the metal stripping section, in contrast to the othersections, the flow of the organic and aqueous phases is cocurrent. Theorganic extractant phase 36 from the wash section is contacted with asoluble calcium phase 52 comprised, for example, of from about 2% toabout 30% by weight calcium nitrate with the remainder being water. Thetemperature of the calcium solution phase 52 can be from about 20° C. toabout 77° C., preferably from about 60° C. to about 71° C. Thistemperature can be achieved by utilizing the heat of dilution ofsulfuric acid with water and with hot steam condensate.

On the FIGURE, the treatment with soluble calcium and the subsequentneutralizations are shown as one in two stages, in vessels 51, 53 and54, 55. If no more than about 90% removal of magnesium and aluminumimpurities is desired, one section (e.g., 54,155) can be eliminated.

In the FIGURE, the aqueous calcium phase 52 is pumped into mixer-settlercell 51 and mixed with the partially stripped organic extractant phase36 coming from mixer-settler cell (or 35, if the wash is omitted). Thetemperature of the resulting mixture is from about 50° C. to about 65°C. After mixing, the mixture of the organic and the aqueous phase isallowed to settle and separate into discrete phases. The immiscibilityof the two phases and the difference in the specific gravities of thephases aids separation. After separation, the organic extractant phase36 is pumped to a second mixer settler cell 54, where it is againcontacted with a soluble calcium phase.

Upon leaving the mixer-settler cell 51, the soluble calcium phase ispassed through a neutralization stage where it is contacted with a base(e.g., lime) to precipitate solids comprising compounds of calcium,magnesium, aluminum, phosphorous (i.e. "P₂ O₅ " values) and sulfate(because commercial wet-process phosphoric acid usually contains in therange of about 0.5 to about 5% sulfate ions some of which is picked upby the organic extractant).

These solids are low in fluorine and are useful as an animal feedsupplement (e.g., for poultry feeding). In the FIGURE the solids areidentified as "solids containing P₂ O₅, Al, Mg".

The neutralized soluble calcium solution ("calcium nitrate solution" inthe FIGURE) is perferably recycled to the mixer-settler 51.Alternatively, (not shown) the neutralized soluble calcium solutioncould be passed to the second mixer settler cell 54.

Similarly, in the FIGURE, the organic extractant phase 36 is passedthrough a second mixer-settler 54 where it is again contacted with asoluble calcium solution 52.

The organic extractant phase 36 from the mixer-settler 54 (or if asingle calcium contact, 51) is contacted in vessel 60 with a mineralacid stripping phase 61 comprised of from about 2% to about 30% byweight sulfuric acid, with the remainder being water. The temperature ofthe mineral acid stripping phase 61 is from about 20° C. to about 77°C., preferably from 60° C. to about 71° C. This temperature is achievedby utilizing the heat of dilution of sulfuric acid with water and withhot steam condensate.

The mineral acid stripping phase 61 is pumped into mixer-settler cell 60and mixed with the partially stripped organic extractant phase 36 comingfrom mixer-settler cell 54, the temperature of the resulting mixture isfrom about 50° C. to about 65° C. After mixing, the mixture of theorganic and the aqueous phase is allowed to settle and separate intodiscrete phases. The immiscibility of the two phases and the differencein the specific gravities of the phases aids separation.

After separation, the barren organic extractant phase 34 is pumped toorganic extractant phase storage (not shown). The mineral acid strippingphase 61a on exiting cell 60 is termed the "waste mineral acid strippingphase". The waste mineral acid stripping phase on exiting themixer-settler cell 60 is reduced in strength to between about 15% to 20%by weight H₂ SO₄. The waste mineral acid stripping phase can bepartially combined with mineral acid stripping phase 61 (not shown) tofurther utilize unused sulfuric acid values in said waste mineral acidstripping phase, or it can be sent to neutralization stage (53 or 55) orto a disposal area (not shown) where the sulfuric acid values andimpurities can be precipitated by the addition of lime or the like towaste mineral acid stripping phase 61a. The waste mineral acid can bepurified by electro-dialysis.

When the organic extractant which is predominantly loaded by calcium ionis regenerated by sulfuric acid, calcium sulfate forms and precipitatesin the aqueous phase. The aqueous phase is easily filterable and thesulfuric acid can be recycled. Since the removal of calcium is dictatedby the low solubility of gypsum, considerably higher strengths ofsulfuric acid can be used than with the prior art processes.

EXAMPLE 1

A stock solution of the organic extractant phase was prepared bydissolving dinonylnaphthalenesulfonic acid (DNSA) anddi-(2-etheylhexyl)phosphoric acid (DEHPA) in mineral spirits containingisodecanol. The composition by weight was as follows:

    ______________________________________                                               DNSA            18%                                                           DEHPA           6%                                                            Mineral Spirits 74%                                                           Isodecanol      2%                                                     ______________________________________                                    

The organic extractant phase, as described above, was contacted withhemihydrate process phosphoric acid containing 41% by weight P₂ O₅ at aphase volume ratio of 6/1 (organic/phosphoric acid) at 54° C. After thephases separated, the organic exxtractant phase was centrifuged toremove any entrained phosphoric acid phase therefrom to produce ametal-loaded organic extractant phase. The loaded extractant was washedwith dilute (1% P₂ O₅) aqueous phosphoric acid to remove P₂ O₅ valuesfrom the loaded extractant phase and produce a washed metal-loadedextractant (hereinafter, sometimes, "WLE"). The ionic metallic impuritycontent and the P₂ O₅ content of the washed organic extractant phase wasdetermined in the following manner. An aliquot of the organic phase wasstripped with equal volumes of 2 N HCl; the total volume of the HCl usedwas equal to the volume of the organic extractant phase analyzed. Theionic metallic impurities were analyzed by means of atomic absorptionspectrophotometric methods. The P₂ O₅ (% by weight) was analyzed by theofficial method of AOAC Method 12, page 13, 11th Edition, 1960*.

The WLE analysis was 0.93 g/l Ca, 0.72 g/l Mg.

1. Substitution step: 800 ml WLE was shaken with 200 ml 5% Ca(NO₃)₂solution for 5 minutes. Twenty minutes' settling was allowed. Theaqueous stream was designated as Sample 2. Both WLE and substituted WLE(henceforth called CaLE or calcium loaded extractant) was shaken withequal volume of 2 N HCl. The aqueous from these operations constitutedSample 1 and Sample 3, respectively, representing calcium and magnesiumion content in the organic phase.

2. One part of CaLE, 300 ml was shaken with 300 ml 25% H₂ SO₄ (O/A=1).The aqueous phase after the exchange was filtered for gypsum and theliquor was designated as Sample 4. The sample prepared from theregenerated organic (2 N HCl strip) was designated as Sample 5.

3. To test the recyclability of the aqueous stream from strip 1 step,another 300 ml of CaLE was stripped with the aqueous stream from c,after the strength was brought back to 25% H₂ SO₄. The sample preparedfrom the regenerated strip from this step was called Sample 7. Theaqueous stream constituted Sample 6.

4. Aqueous stream from substitution step (Sample 2) was limed until pHrose to 10.6. The precipitate containing magnesium was filtered off.

5. These samples analyzed for calcium, magnesium, iron, aluminum, etc.

The results of the above tests were as follows:

1. Using a 5% Ca (NO₃)₂ solution, 43% of magnesium in the WLE wastransferred to the aqueous phase. Also 3.45 atoms of calcium formedcalcium sulfonate per atom of calcium replacing magnesium in theextractant. A preliminary test previous to this test showed a 92%magnesium removal using a more concentrated calcium chloride solution.Thus a stronger solution of calcium nitrate will be needed to increasemagnesium removal at this stage.

Analyses of the streams for calcium, magnesium and aluminum adequatelyrepresented material balance as shown in the following diagram.

    __________________________________________________________________________     ##STR1##                                                                     Total Ca in = 3.26 gm   Total Ca out = 2.9 gm                                 Total Mg in = 0.58      Total Mg out = 0.58                                   Total Al in = 0.08      Total Al out = 0.3                                    __________________________________________________________________________

2. Waste Ca (NO₃)₂ Stream Liming:

Analysis of the limed stream showed that 95.5% of magnesium in thestream was removed by liming. Aluminum and iron were almostquantitatively removed.

3. Stripping with H₂ SO₄ :

Both stripping operations produced satisfactory regenerated extractantscontaining comparable calcium and magnesium (within experimental error).The recyclability of the sulfuric acid strip liquor is establishedbeyond any doubt. The material balance closure was also satisfactory.This is shown in Table 1.

                  TABLE 1                                                         ______________________________________                                        Material Balance in Stripping                                                                Stripping 1                                                                             Stripping 2                                          ______________________________________                                        Total calcium in 0.81 gm     0.91 gm                                          Total calcium out                                                                              0.11 gm     0.12 gm                                          Calcium precipitated                                                                           0.69 gm     0.80 gm                                          Total magnesium in                                                                             0.12 gm     0.26 gm                                          Total magnesium out                                                                            0.14 gm     0.21 gm                                          Total aluminum in                                                                              0.03 gm     0.056 gm                                         Total aluminum out                                                                             0.03 gm     0.04 gm                                          Concentration of Ca in BE                                                                      0.03 gm/l   0.04 gm/l                                        Concentration of Mg in BE                                                                      0.02 gm/l   0.02 gm/l                                        ______________________________________                                    

EXAMPLE 2

70 ml of a 10% Ca(NO₃)₂ solution was shaken in a separating funnel with170 ml of WLE (O/A=2.43) at room temperature for 5 minutes. After phasedisengagement the aqueous stream was subjected to sequential pHadjustments with lime. At pH 7.7 phosphates precipitated and thefiltrate was further treated with lime to attain a pH of 10.5, whenmagnesium compounds precipitated. The experimental scheme is shown belowwith sample identifications. The analysis for components of thesesamples are presented in Table 2.

                  TABLE 2                                                         ______________________________________                                        ANALYSIS OF SAMPLES                                                           Sample #  Mg. g/l   Ca. g/l  Al. g/l                                                                              P.sub.2 O.sub.5. g/l                      ______________________________________                                        1         0.84      15.9     0.20   5.65                                      2         0.51      18.6     0.00   0.01                                      3         0.01      19.2     0.00   0.01                                      ______________________________________                                    

This Example 2 establishes that:

Controlled liming of the calcium nitrate solution from the substitutionstep of a slightly alkaline pH (e.g. 7.7) removed entrained phosphatequantitatively. This recovery of the phosphate improve the economics.The recovered P₂ O₅ was entrained in WLE.

A second stage liming to adjust the pH of the supernatent liquid toabout 10.5 removed magnesium from solution without precipitating thecalcium. Thus, the calcium nitrate stream can be regenerated andrecycled.

EXAMPLE 3

A continuous pilot plant run of the process, using calcium nitratestripping was conducted. This run (Run 1) lasted for six hours. Theoperating temperature was 95° F. Washed loaded extractant (WLE), similarto that in Example 1, was contacted in two stages with 10% calciumnitrate solution. The calcium loaded extractant was contacted with 25%H₂ SO₄ in the third contactor.

The three phase (CaSO₄ 2H₂ O, organic and aqueous sulfuric acid) slurrywas then fed to a thickener. Solid calcium sulfate and sulfuric acidwere withdrawn from the bottom of the thickener, while the regeneratedorganic (barren extractant) was withdrawn from a weir at the top. Thegypsum filtration characteristics were excellent.

The conditions of the run were as follows:

    ______________________________________                                        extractant             40 cc/min                                              Ca(NO.sub.3).sub.2, 1st stage                                                                        20 cc/min                                              Ca(NO.sub.3).sub.2, 2nd stage                                                                        20 cc/min                                              25% H.sub.2 SO.sub.4   20 cc/min                                              O/A ratio              2/1                                                    Residence time in mixers                                                                              5 minutes                                             ______________________________________                                    

Samples were taken from various locations and analyzed for theefficiency of the exchange and regeneration of the extractant. Resultsare shown in Table 3.

                  TABLE 3.                                                        ______________________________________                                        Strip Results of Run 1                                                                   Ca,    Mg,     Al,   Fe,   NO.sub.3,                                                                          P.sub.2 O.sub.5,                   Sample     g/l    g/l     g/l   g/l   g/l  g/l                                ______________________________________                                        Washed loaded                                                                 extractant 0.24   0.95    0.15  0.21  0.29                                    Barren extractant                                                                        0.17   0.013   <.02  0.117 0.11 0.063                              Calcium nitrate                                                               effluent from                                                                 from Stage 1                                                                             14.1   1.375   0.205 0.045 46.5 5.3                                Calcium nitrate                                                               effluent from                                                                 from Stage 2                                                                             19.85  0.37    0.065 0.005 64.0 1.15                               ______________________________________                                    

This run showed a 92% efficiency in calcium-magnesium exchange in thecalcium nitrate treatment stages. Quality of the barren extractant wasgood as shown in the table. Two other runs (Run #3 and Run #4) were madein which the regeneration of the calcium nitrate effluent and thesulfuric acid stream was done and the streams recycled.

What is claimed is:
 1. In a process for purifying an aqueous phosphoricacid containing P₂ O₅ values and dissolved ionic impurities comprisingmagnesium, iron and aluminum ions, wherein (a) said phosphoric acid iscontacted with an organic extractant containing a water-immiscibleorganic sulfonic acid to form: a loaded organic extractant containingthe organic sulfonic acid, extracted ionic metallic impurities andextracted P₂ O₅ values, and a purified aqueous phosphoric acid; (b) thepurified aqueous phosphoric acid is separated from the loaded organicextractant to provide a separated organic extractant; (c) said separatedorganic extractant is contacted with a solution containing calcium ions,under conditions whereby at least some of said extracted ionic metallicimpurities are replaced by calcium and are transferred to said solution,thereby producing the calcium form of the extractant; and (d) saidsolution is separated from the contacted extractant comprising saidcalcium form of the extractant; the improvement comprising(i) contactingsaid separated extractant, from step (c) comprising said calcium form ofthe extractant, with sulfuric acid to convert at least some of saidcalcium form of the extractant to the hydrogen or acid form therebyproducing an acidified extractant, used sulfuric acid and solid calciumsulfate; (ii) separating said solid calcium sulfate, said acidifiedextractant and the used sulfuric acid; (iii) recycling said usedsulfuric acid to step (i) for further contact with more loadedextractant; and (iv) recycling said acidified extractant to step (a) forfurther contact with more phosphoric acid.
 2. The process of claim 1wherein, in step (iii), additional, stronger sulfuric acid is added tothe used sulfuric acid which is recycled to step (i).
 3. The process ofclaim 1 wherein said separated organic extractant is washed with wateror dilute phosphoric acid prior to said contact of step (c).
 4. Theprocess of claim 1 wherein said separated solution in step (d) iscontacted with a sufficient amount of a base to cause a solid to form.5. The process of claim 4 wherein said base contains calcium ions andsaid solid contains calcium phosphate.
 6. The process of claim 5 whereinsaid base contains lime.
 7. In a process for purifying an aqueousphosphoric acid containing from about 1% to about 55% P₂ O₅ by weightand dissolved ionic impurities comprising magnesium, iron and aluminumions, wherein (a) said phosphoric acid is contacted with an organicextractant containing a water-immiscible organic sulfonic acid to form:an organic extractant containing the organic sulfonic acid, extractedionic metallic impurities and extracted P₂ O₅ values, and a purifiedaqueous phosphoric acid; (b) the purified aqueous phosphoric acid isseparated from the organic extractant to provide a separated organicextractant; (c) said separated organic extractant is contacted with asolution containing aqueous calcium nitrate, under conditions whereby atleast some of said extracted ionic metallic impurities are replaced bycalcium and are transferred to said solution phase, thereby producingthe calcium form of the extractant; (d) said solution is separated fromthe contacted extractant; (e) said separated organic extractant isacidified with sulfuric acid to produce solid calcium sulfate andconvert at least some of the calcium form of the extractant to thehydrogen or acid form and causing solid calcium sulfate to form; (f)separating the acidified extractant from the used sulfuric acid and thecalcium sulfate; and (g) recycling at least part of said acidifiedextractant from step (f) to step (a) for contact with said phosphoricacid, the improvement comprising separating calcium sulfate from theused sulfuric acid in step (f) and recycling the separated, usedsulfuric acid to step (e).
 8. The process of claim 7 wherein theseparated solution from step (d) is contacted with sufficient lime tocause calcium phosphate to precipitate.
 9. The process of claim 7wherein prior to said contact of step (c) said separated organicextractant is washed with water or dilute phosphoric acid containing inthe range of about 10 grams per liter to about 300 grams per liter of P₂O₅.
 10. The process of claim 7 wherein the separated solution from step(d) is contacted with sufficient lime to cause solids to form whichcontain compounds of magnesium.
 11. The process of claim 10 wherein saidsolids are recovered from said solution and are suitable for use as aplant or animal nutrient.